Combination of dy alumino-silicate catalyst and hydrogenation catalyst



Oct. 13, 1970 R. D. BUSHICK 3,534,115 COMBINATION OF D ALUMINO-SILICATE CATALYST AND HYDROGENATION CATALYST Filed Aug. 2. 1968 r 5 Sheets-Sheet 1 FEED INERT FEED PARAFFIN GAS OLEFIN 3 FEED SECTION l7 Ml FIG.I

INVENTOR RONALD D. BUSHICK I l"! 'v k4 lz l l ny y/l/Jr h ATTORNEY Filed Aug. 2. 1968 Oct. 13, 1970 R. D. BUSHICK 3,534,115

COMBINATION OF Dy ALUMINO'SILICATE CATALYST AND HYDROGENATION CATALYST 5 Sheets-Sheet 2 REACTOR SECTION HIGH LOW P P 57 LOW U PRESSURE ADJUVANTS 62 6 HCI H20, ETC) 37 34 v 33 r\ REACTION FEED PRODUCT T0 35 REACTOR 39 i 36 610' v 53 I v 4| 40 HIGH PRESSURE 43 INERT GAS 4 INVENTOR RONALD D. BUSHICK ATTORNEY Oct, 13, 1970 R. D. BUSHICK 3,534,115

COMBINATION OE DY ALUMINO-SILICATE CATALYST AND HYDROGENATION CATALYST Filed Aug. 2, 1968 '5 Sheets-Sheet :5

PRODUCT RECOVERY SECTION REACTION PRODUCT LIQUID PRODUCT (ALKYLATE) FI'G.3

INVENTOR RONALD D. BUSHICK United States Patent Ofiice I 3,534,115 Patented Oct. 13, 1970 3,534,115 COMBINATION OF Dy ALUMINO-SILICATE US. Cl. 260-668 15 Claims ABSTRACT OF THE DISCLOSURE A process for converting hydrocarbons comprises contacting a hydrocarbonaceous feed in a conversion zone ('with from -5000 p.s.i.,g. hydrogen) at an elevated temperature with a novel composition comprising an acidic Dy alumino-silicate catalyst and from 0.5 to 25% by weight of a hydrogenation catalyst, and recovering an upgraded hydrocarbon conversion product. The Dy alumino-silicate catalyst contains less than one alkali metal cation and at least one divalent or trivalent metal, metal oxide or metal hydroxide for every 12 atoms of aluminum in the alumino-silicate zeolite framework.

The hydrogenation catalyst can be platinum, palladium, nickel, nickel oxide, nickel sulfide, molybdenum oxide, molybdenum sulfite, cobalt oxide and palladium oxide. The hydrogenation catalyst can be physically admixed with the Dy catalyst, or incorporated therein by salt impregnation or ion exchange.

The novel catalyst combination is especially useful for the conversion of polycyclic aromatic hydrocarbons, such as converting s-octahydroanthracene to such compounds as Tetralin, 1-cycloheXyI-Z-phenylethane, asymmetrical octahydroanthracene, asymmetrical octahydrophenanthrene, 1,2,3,4 tetrahydroanthracene, 1,2,3,4 tetrahydrophenanthrene, anthracene, phenanthrene, and s-octahydrophenanthrene.

CROSS REFERENCES TO RELATED APPLICATIONS Dy-containing zeolites which can be utilized in the claimed catalyst combination and in the claimed process have been described in copending application Ser. No. 590,225, filed Oct. 28, 1966, now US. 3,396,203 of Ronald D. Bushick entitled Alumino-Silicate Catalyzed Reaction of Polycyclic Aromatic Hydrocarbons of which this application is a continuation-in-part and in copending application Ser. No. 750,432, filed Aug. 2, 1968, of Ronald D. Bushick entitled Process for Producing Sym- Octahydroanthracene From Sym Octahydrophenanthrene, and in copending application Ser. No. 581,129, filed Aug. 25, 1966 of Francis William Kirsch, David S. Barmby and John D. Potts entitled Process for Parafi'in- Olefin Alkylation and in copending application Ser. No. 716,190, filed Mar. 26, 1968 of Francis William Kirsch, David S. Barmby and John D. Potts entitled Process for Paraffin-Olefin Alkylation and in copending application Ser. No. 749,714, filed of even date with the present application of Francis William Kirsch, David S. Barmby and John D. Potts entitled Dy Zeolite and Hydrocarbon Conversion Process With Dy Zeolite Catalyst, all of these being assigned to Sun Oil Company. The disclosure of all of the above-cited applications is hereby incorporated in the present application.

BACKGROUND OF THE INVENTION Although it is known to utilize crystalline aluminosilicate zeolites containing cations of lanthanum, cerium or of certain rare earth salt mixtures as hydrocarbon con version catalysts (e.g., see US. 3,140,249 and US. 3,210,- 267), the art has failed to realize that incorporation of substantial quantities of dysprosium cations in an amorphous or crystalline alumino-silicate zeolite can be used to produce an adsorbent for aromatic hydrocarbons or a catalyst which is especially useful for hydrocarbon conversion reactions, and particularly conversions involving carbonium-ion mechanisms. Similarly, the art has failed to appreciate that a combination of a Dy alumino-silicate catalyst with a hydrogenation catalyst can be especially useful for conversions involving catalytic contacting of hydrocarbons in the presence of hydrogen, such as aromatization of naphthenes or olefins, cyclizations, reforming, hydrocracking and hydroisomerization.

BRIEF SUMMARY OF THE INVENTION Hydrocarbon conversion reactions, such as cracking, dehydrogenation, reforming, alkylation, dealkylation, cyclization and isomerization can be effected by contacting a hydrocarbon feed with a catalyst comprising an aluminosilicate zeolite containing cations of dysprosium, such as Dy+ Dy(OH)+ and Also effective as catalysts in such processes are novel catalysts comprising dysprosium-containing zeolites which also contain magnesium cations, aluminum cations, silver cations, nickel cations, zinc cations, cerium cations, lanthanum cations, gadolinium cations, cations of the hydroxides or oxides of these metals or mixtures of two or more of such cations.

The preferred zeolite catalyst is crystalline and capable of adsorbing benzene, has an atomic ratio Al/Si of 0.65 to 0.35 and contains at least one cation for every 8 atoms of aluminum in the aluminosilicate framework. The zeolite can also be utilized as an adsorbent, as for separating aromatic hydrocarbons from less polar compounds. In conversions involving oxidative regeneration of this catalyst (or adsorbent), crystallinity can decrease, usually accompanied by a decrease in activity and/or selectivity. The resulting, more amorphous, zeolite can be effective as a catalyst, particularly at conversion temperatures which are greater than those required for the corresponding conversion with an equal weight of crystalline zeolite.

An especially useful hydrocarbon conversion reaction is the hydroisomerization of the C -C parafiin hydrocarbons which are capable of conversion to move highly branched isomers, in order to obtain a more highly branched product with improved octane ratings. Another especially useful conversion is the alkylation of C C aromatic (or C -C alkyl aromatic) hydrocarbons, or C -C paraffin hydrocarbons, with olefin hydrocarbons. Another especially useful conversion is the isomerization of such polycyclic aromatic hydrocarbons as s-octahydrophenanthrene (s-OHP) to produce s-octahydroanthracene (s-OHA) and/or such aromatic hydrocarbons as l-cyclohexyl-Z-phenyl ethane, asymmetrical octahydrophenanthrene, 1,2,3,4 tetrahydroanthracene, 1,2,3,4-tetrahydrophenanthrene, anthracene, phenanthrene and Tetralin. It is sometimes also advantageous to conduct such aromatic isomerization reactions in the presence of from 5 to 1500 p.s.i. of hydrogen.

FURTHER DESCRIPTION OF THE INVENTION Parafiin-olefin alkylation activity can be a measure of the acid activity of a catalyst, and, therefore, indicative of the ability of the catalyst to catalyze typical carboniumion reactions such as cracking, dealkylation, aromatic alkylation, polymerization, isomerization, etc. That is, a

necessary, but not sufficient, condition for the catalytic alkylation of isobutane with a C -C monoolefin is that the catalyst must have the ability to generate carboniumions (one further requirement for parafiin-olefin alkyla tion is good hydride transfer ability). By utilizing the liquid phase alkylation of i-butane with butene-2 as a test reaction, it has been found that a substantially anhydrous DyHY catalyst, prepared by activation of a crystalline DyNH Y zeolite (obtained by Dy-cation exchange of highly amonium-exchanged sodium Y zeolite) is an efiective catalyst for hydrocarbon conversion reactions, especially those which are considered by the art to involve a carbonium-ion mechanism.

In the Dy alumino-silicate catalyst, at least 25 and, preferably, at least 40% of the electronegativity associated with the alumino-silicate framework is satisfied by cations of dysprosium or of its oxides or hydroxides. When the Dy catalyst contains less than one alkali metal cation (e.g., Na+) for every 4 aluminum atoms in the aluminosilicate framework, the catalyst is especially useful for such hydrocarbon conversion reactions as isomerizing polycyclic aromatic hydrocarbons, paraflin-olefin alkylation and cracking of gas oil. Preferably, the aluminosilicate zeolite is crystalline and is chemically characterized by the empirical formula where x, y and z are integers, the ratio x/y being from 1.0 to 0.2 and where M is chosen from at least one of the following groups:

(1) at least one Dy+ cation for every 12 atoms of aluminum in the alumino-silicate framework of said zeolite;

(2) at least one cation of Dy(OH)+ for every 8 atoms of aluminum in the alumino-silicate framework of said zeolite;

(3) at least one cation of for every 4 atoms of aluminum in the alumino-silicate framework of said zeolite;

(4) a combination of the members of at least two of the above groups;

and wherein the balance of the cations necessary for electronic equivalency comprises H+ or cations of metals, metal oxides, or metal hydroxides and wherein there is less than one alkali metal cation for every four atoms of aluminum in the alumino-silicate zeolite, more preferably, less than one alkali metal cation for every ten atoms of aluminum.

The Dy zeolite can contain as such addition cations, the cations of magnesium, aluminum, silver, nickel, zinc, cerium, gadolinium, lanthanum and mixtures of these cations. In such catalysts it is preferred that at least one such cation is present for every 20 atoms of aluminum in the alumino-silicate framework of said zeolite.

For most hydrocarbon conversions, the ratio x/z in the empirical formula of the zeolite should be in the range of 0.25 to 2. If excess water is present, the zeolite should be activated by heating according to the procedure disclosed in the aforementioned applications of Kirsch, Barmby and Potts. If the zeolite is deficient in bound water, such water can be added, as by exposure to steam in air or nitrogen.

As used herein, the term framework, in reference to the alumino-silicate portion of the zeolite (which can be crystalline or amorphous), excludes those aluminum ions which are in exchange positions and which are neutralizing some of the negative charge associated with the aluminum atoms in the alumino-silicate tetrahedra of the zeolite. Note that aluminum in the alumino-silicate framework can be either trigonal or tetrahedral.

For such reactions as reforming, aromatization, hydro gen transfer, hydrocracking and hydroisomerization, it is p eferred that the catalyst have incorporated therewith from 0.05 to 25% (more preferably, 0.05 to 5%) of a hydrogenation catalyst component containing a hydrogenactive metal such as platinum, palladium, rhodium, rhenium, ruthenium, molybdenum, cobalt or nickel (or a chemical compound, as an oxide or sulfide, of such a metal). The hydrogen-active metal can also be incorporated on a carrier (as alpha-alumina, microporous silica, conventional amorphous silica-alumina cracking catalyst, or acid-exchanged clays, such as montmorillonites or kaolin). When the hydrogen-active metal component (or a chemical compound of the metal) is so incorporated on a carrier, it is preferred that the Dy catalyst be physically admixed therewith.

When the hydrocarbon conversion involves cyclization and/or aromatization, as with a feed of n-pentene, nhexene, n-heptene or 1,4-dimethylnaphthalene, the cyclization conditions comprise a temperature in the range of 350850 F. and a pressure in the range of 0750 p.s.i.g., preferably with the reactants maintained in the vapor or trickle phase. For aromatization and/or cyclization of a cracked naphtha, temperature in the range of 240600 F. is preferred at atmospheric pressure. For a hydrogen transfer reaction, to produce aromatics from naphthenes, a temperature in the range of 300500 F. at atmospheric pressure is preferred, as when cyclohexane and propylene are the feed hydrocarbons and the products are benzene plus propane.

For double-bond isomerization, such as for the conversion of 2-ethyl-l-butene to cis and trans 3-methyl-2- pentene, or the conversion of pentene-l to pentene-2, a temperature in the range of 70400 F. and pressures from 0-75 p.s.i.g. are preferred, with the lower temperatures and higher pressures most preferred in order to reduce cracking.

For isomerization and/or transalkylation of alkyl benzenes, such as converting meta-xylene to ortho and paraxylene, the hydrocarbon reactant can be either in liquid or vapor phase at a temperature in the range above about 60 C. and below cracking temperature. The preferred temperature range for xylene isomerization is -350 C. and preferably in the presence of added hydrogen (e.g., 5-75 p.s.i.).

When the primary conversion reaction is cracking, a temperature in the range of 800-1l00 F. is preferred for a gas oil feed, preferably at atmospheric or slightly elevated pressure, although pressures as low as 1 mm Hg and as high as 1200 p.s.i.g. can be utilized in such cracking reactions. When the predominant reaction is hydrocracking, our preferred hydrogen-active metal is selected from Group VIb, VIII, and more preferably comprises Ni, Mo, Co, Pd or Pt. The preferred hydrogen pressure is in the range of 500-5000 p.s.i. at conversion temperatures from 6501l00 F.

For parafiin-olefin alkylation, the preferred process conditions with a C C feed olefin are those of the aforementioned patent application of Kirsch, Barmby and Potts. Generally, these involve (a) contacting C -C monoolefin with C -C isoparaffin and with a substantially anhydrous Dy zeolite catalyst at a temperature below the critical temperature of the lowest boiling hydrocarbon reactant and at a pressure such that each of the reactants is in liquid phase, and

(b)stopping such contacting after substantial alkylation has occurred but before the weight rate of production of unsaturated hydrocarbon becomes greater than the weight rate of production of saturated hydrocarbon.

Preferably, the feed olefin and feed paraffin are admixed prior to contact with the catalyst and the concentration of unreacted olefin is kept sufliciently low that predominantly saturated paraflin-olefin alkylation products are obtained rather than unsaturated products. This concentration is preferably less than seven, more preferably less than 12 mole percent, based on the total paraflin content of the reaction mixture. Also preferred is the use of a halide adjuvant (as HCl, CCL, and the C C monochloro parafiins) containing bromine, chlorine, or fluorine, to increase the yield of liquid paraffin based on the olefin charged. Also preferred is a temperature in the range of 25-120 C. and a mean residence time of the reaction mixture with the catalyst in the range of 0.05 to 0.5 hour per gram of catalyst per gram of hydrocarbon in the reaction mixture. When the feed olefin comprises ethylene, conditions shown in US. 3,251,902 can be used to produce a liquid product; however, this liquid product is generally less preferred as a component of gasoline than are the highly branched liquid parafiin hydrocarbons which are produced by the aforementioned processes of the Kirsch, Barmby and Potts applications.

For the isomerization of such polycyclic aromatic hydrocarbons as s-OHA to its isomer s-OHP, or a s-OHP to its isomer s-OI-LA, the preferred conditions include a temperature above 80 C. and below cracking temperature and are shown in the aforementioned .U.S. Pat. 3,- 396,203 of Ronald D. Bushick of Aug. 6, 1968. This Bushick patent also shows the preparation of a novel composition comprising an acidic Dy alumino-silicate catalyst and from 0.5 to 5% of a hydrogenation catalyst. Preferably, the hydrogenation catalyst is selected from the group consisting of platinum, palladium, nickel, nickel oxide, nickel sulfide, molybdenum oxide, molybdenum sulfide, cobalt oxide, palladium oxide and mixtures thereof. The hydrogenation catalyst can be physically admixed with the acidic alumino-silicate, or have been incorporated into the alumino-silicate by salt impregnation or by ion exchange. When the salt has been introduced into the alumino-silicate catalyst by ion exchange, it is preferred that the hydrogenation catalyst be reduced, as with hydrogen, prior to contact of the catalyst with the hydrocarbon feed. Also preferred is a process for the isomerization of polycyclic aromatic hydrocarbons, such as s-OHA or s-OHP, wherein the Dy catalyst/hydrogenation catalyst combination and from 25-1000 p.s.i.g. of hydrogen are present in the reactor. The added hydrogen aids in maintain the activity of the isomerization catalyst combination, and can be recycled at rates up to 10,000 s.c.f./ bbl. of feed. The LHSV is preferably in the range of 0.255.0 volumes of feed per volume of catalyst per hour. Such added hydrogen is also sometimes advantageous when no such hydrogenation catalyst is present with the Dy catalyst.

In any of the above-noted reactions, if the catalyst activity appreciably decreases during the course of the reaction, the catalyst may be separated from the hydrocarbon reactants and regenerated, as by purging with ammonia or burning in air. After such purging or burning, water can be added to the catalyst, as by exposure to hydrogen or to steam in air or nitrogen. When a hydrogen-active metal is incorporated into the zeolite catalyst, it is sometimes advantageous to reduce the regenerated combination with hydrogen, preferably at 250 to 800 F., prior to introduction of the hydrocarbonaceous feed.

When the primary hydrocarbon conversion is a hydroisomerization or a reforming reaction, the preferred conditions include a hydrogen pressure of at least 25 p.s.i.g. an temperatures from 500 to 700 F., although the conversion can be effected in the range of 225100 F., at total pressures in the range of -5000 p.s.i.g. and hydrogen pressures in the range of 0.5-4000 p.s.i.g. For the hydroisomerization of C -C parafiins, the preferred catalyst combination will contain from 0.1 to 2 percent of Pt, Pd or Re (or a mixture thereof) or from 1 to of Ni.

Typical feeds and reaction conditions which are effective when utilizing the Dy catalyst, particularly when combined with a hydrogenation catalyst, for hydroisomerization or reforming are those in the following US. patents: 2,834,439; 2,970,968; 2,971,904; 2,983,670;

3,190,939; 3,197,398; 3,201,356; and 3,236,762.

In conversion processes utilizing the Dy catalyst, whether alone or in combination with a hydrogenation catalyst, halide adjuvants containing chlorine, fluorine or bromine can frequently be used to increase the degree of conversion. Preferred adjuvants include CCl HCl, AlBr BF HF and the C -C chlorohydrocarbons.

DESCRIPTION OF THE DRAWINGS FIGS. 1, 2 and 3 illustrate the three basic sections which comprise a continuous stirred reactor system which can be used to produce an olefin-parafi'in alkylate using an acidic Dy-zeolite catalyst and the process of the aforementioned U.S. patent applications Ser. No. 581,129 and Ser. No. 716,190. FIG. 1 illustrates the feed section, FIG. 2, the reactor section and FIG. 3, the recovery section.

In FIG. 1, the valving arrangement at the top of the mixing vessels 17 and 18 allows feed parafiin 1 or feed olefin 3 to be placed in either vessel 17 or vessel 18 from either the top or the bottom of the vessel. For example, parafiin can be introduced through the bottom of vessel 17 by closing valves 9, 12, 19 and 23 and opening valves 4, 8 and 21. Then, feed olefin 3 is transported to vessel 17 by closing valves 4, 11, 12, 14, 19 and 23 and opening valves 7, 8, 9, 10 and 21. Alternately, the mixing of the incoming feed olefin and feed parafiin can be effected by means of an inline mixer; however, for precise control of the reactant proportions and to insure intimate admixing of parafiin and olefin, we prefer that a substantial amount of parafiin admixed with olefin be maintained in a stirred mixing vessel as vessels 17 and 18.

Similarly, by sequencing the position of the valves, the feed paraffin and the feed olefin can be introduced in anydesired pattern. One sequence of placing feed parafiin and feed olefin in vessel 17 is to allow feed parafiin to enter vessel 17 to a level a. Sufiicient feed olefin is then brought into vessel 17 to produce a volume of parafiin-olefin admixture represented by level b. The remainder of the required parafiin feed is added to vessel 17 until the level of the total feed mixture is at c. Such a sequence of parafiin-olefin-parafiin addition allows for better internal mixing of the reactants in vessel 17 (in addition, uniform mixing is insured by mixing devices 15, 16, such as, turbine blade rotary mixers). We have also found that additional mixing can be accomplished by bringing some or all of the inert gas into the bottom of the mixing vessel as through valve 21, rather than into the top of the vessel, as through valve 12.

It is generally preferable to introduce feed components to the mixing vessels in a number of alternate portions (unless the feed components are simultaneously proportioned into an inline mixer) to insure uniform mixing. Similarly, sequencing of valves can be used to fill vessel 18 while the mixture in vessel 17 is being fed to the reactor 43 (of FIG. 2). In order to use vessel 17 regardless of whether vessel 18 is being filled or not, pressure, as by inert gas 2 (e.g., nitrogen) is imposed upon the liquid in vessel 17, as by closing valves 12, 19 and 21 and opening valve 5. Normally, the nitrogen head is allowed to build up until the pressure in the mixing vessel is about 50 p.s.i. less than the pressure in the reactor 43.

In order to allow a mixed parafiin-olefin feed to enter the reactor 43, the nitrogen head is imposed upon vessel 17 and valve 19 is opened. A constant head pressure on vessel 17 allows the pump 31 to pump the mixed feed through line 33 to the reactor at a constant rate.

When valve 19 is open, the feed mixture passes through a microfilter 24 (which protects the pump and meters from damage caused by foreign particles), then through a high pressure rotometer 29 (which serves as a flow indicator). The feed can then enter the pump 31 when valve 30 is open and, when valve 32 is open and 34 closed, the feed is pumped into the reactor 43. In the event of a pump failure, valve and valve 32 may be closed, needle valve 34 opened and the nitrogen head increased sufficiently to allow the feed to flow through line 35 to line 33 and, thence, to the reactor 43.

FIG. 2 illustrates the reactor section, comprising a continuous stirred reactor vessel and the associated lines and valving required for introducing feed, removing reaction products, and for operation of the differential pressure cell which is used for liquid level control. The reactor also contains heat transfer and control means (as a water jacket and electrical heaters, not shown) for maintaining the desired reaction temperature. The paraflin-olefin feed from line 33 enters the reactor 43 through valve 34 and line 35. To insure maximum olefin dilution, we prefer that the liquid feed be allowed to enter the reactor 43 below the reactor liquid level 42 and in the vicinity of the mixing means 36. Halide adjuvants such as CCL; and/or t-butyl chloride can also be metered into the paraflinolefin feed through an additional line, 34a, entering valve 34.

The liquid level is controlled by a differential pressure cell, hereinafter DP cell, having a high pressure section 61 and a low pressure section 60, the differential pressure being in the range of 5-50 inches of Water column. Inert gas 49 enters the DP cell through an inline filter 50 from which it diverges through meter 52 to the high pressure section 61 and through meter 57 to the low pressure section 60. That is, for the high pressure section, valve 51 is open allowing the inert gas stream to flow through the high pressure meter 52 through open valve needle 53 and valve 64 into the high pressure side 61 of the DP cell, then through valve 62 into a line 67 which leads below the liquid level 42 in the reactor. Pressure gauges, 54a and 54b, indicate the pressure in the high pressure side and the lower pressure side, respectively. Other pressure gauges, thermometers, and analytical devices, can be advantageously incorporated into the three sections comprising the apparatus of FIGS. 1, 2 and 3; however, for simplicity such devices are not shown in the figures.

The inert gas can also be diverted through valve 56 to the low pressure meter 57 through needle valve 58 to the low pressure side 60 of the DP cell and then through valve 63 and line 68 to the vapor space above the liquid level 42 in the reactor. In operation, the DP cell senses a differential pressure which is equal to the height of liquid through which the inert gas from the higher pressure side of the cell must travel from the bottom of line 67 (which must be below the liquid level) to the vapor space 39. The difierence between the pressure of the high side 61 and the pressure of the low side 60 of the DP cell is equal to the pressure required to push a bubble of gas through the height of the liquid. Since the DP cell measures the mass of a column of fluid (pressure) and not the volume, its measurement is independent of temperature and, although at a given temperature the actual level of the liquid will vary somewhat, the mass of the volume of liquid above the opening line 67 can be maintained at a constant value regardless of the temperature and pressure of the reactor.

The nitrogen (or other inert gas) which is introduced through the DP cell can be vented through a valve system 40, which can consist of an Annin control valve (a splinetype, highly sensitive metering device) and a block valve ahead of the Annin valve. Preferably, the outlet gas is passed through a stainless steel microfilter (not shown) to prevent inadvertent plugging of the control valve by entrained catalyst. The Annin valve can be actuated by a pressure transmitter (and gas meter) 38 in order to maintain a constant pressure in the vapor space of the reactor.

Catalyst-free reaction mixture is removed from the reactor via line 37 through valve system 69 (which can consist of an Annin valve and a hand-block valve ahead of the Annin valve). The liquid reaction mixture is separated from the suspended catalyst particles by means of a submerged screen 44 and is withdrawn from the reactor through line 37. Although screen plugging is not a frequent occurrence, if plugging occurs the screen can'be back-flushed with nitrogen. This nitrogen back-flush can enter the reactor through line 37 and the excess nitrogen vented through valve system 40 in order that the reactor pressure is maintained constant. This flushing can be effected while the catalyst particles are maintained in suspension and the reaction mixture is maintained in contact With the catalyst particles. In the event that catalyst must be added or removed from the reactor, it may be accomplished through line 48, flush valve 47 and line 46. Similarly, the entire contents of the reactor can be drained through the flush valve 47, if the reaction mixture becomes contaminated or if for any other reason it is desired to drain the reactor contents. For example, if scale builds up on the reactor baflles 45, the reaction mixture can be dumped by opening valve 47, then cleaning materials can be pumped into (and removed from) the reactor through the same valves and lines.

The gases removed via valve 40 can be sent to a gas meter (and pressure transmitter) 38 which can also contain devices for chemical analysis or sampling. For ex ample, the gases so removed can contain HCl, from the halide adjuvant. The HCl concentration in the vapor space is preferably maintained at a constant partial pressure, as by adjusting the quantity of adjuvant which enters the reactor via line 65, valve 66 and line 67. Such adjuvants can also be introduced into the reactor if they are directly added to the paraffin-olefin feed in the mixing vessels 17 and 18. Alkylate yield can be increased and catalyst life prolonged by addition of anhydrous HCl to the reactor at rates in the order of 0.001 to 0.25 mole per mole of olefin charged. Preferably, such HCl addition is regulated such that the concentration of HCl gas in the vapor space 39 above the liquid level 42 in the reactor 43 is maintained at from 0.25 to 25% (more preferably, less than 10%) of the total volume of gas in the vapor space. That is, the partial pressure of the HCl is maintained at from 0.25 to 25 of the total pressure. Thus, the rate of HCl addition depends, primarily, on the volume of the vapor space (which can be adjusted by means of the liquid level control), the temperature, the total pressure and the rate at which the total gases are flowed (e.g., recycled) through the vapor space.

The catalyst-free liquid reaction product (comprising C alkylate, unreacted feed isoparafiin, some C -paraffin product and, usually, a small amount of unreacted feed olefin) which is removed from the reactor via line 37, passes through valve 69 (where the pressure is reduced from reactor pressure to 25 p.s.i.g. or less) and to condenser 72. The condensed liquid and noncondensed gas (e.g., feed isoparafiin) and inert gas (e.g., nitrogen) pass through valve 86 into vessel 92 or, alternately, through valve 81 to vessel 95. We prefer to have two such collecting vessel in order that product can be collected in one vessel while product is removed from the other vessel. The liquid product removed from these vessels can be transported to produce storage tanks or to a means for blending the alkylate with other gasoline components in order to make a blended gasoline product which can be transported to a stabilizer and then to storage area or to tank trucks, etc.

In the product recovery section illustrated in FIG. 3, the liquid produced by condensation of gaseous products in condenser 72 and uncondensed gases pass through line 77 and valve 86 (valve 81 is closed) and enter vessel 92, which is maintained at a temperature and pressure such that liquid alkylate can be removed via valve 96 through line 98 (valve 97 is closed) to tank trucks, a blending area, storage tanks, etc. Uncondensed gases (which consist primarily of unreacted feed isoparaflin) leave vessel 92 via valves 93, and 74 (valves 94, 89 and 73 being closed) and can be passed to means for gas purification and separation 71 or, under some conditions, can be recycled to the reactor or to the mixing vessels via line 76, or via valve 73 and line 70. Minor amounts of the halide promoter which may be present in the reaction product can be removed as by means of an adsorbent which can be between valve 69 and the condenser 72 or at any other appropriate location in the product recovery section. When some or all of the halide promoter is a readily distillable gas such as HCl or methyl chloride, such gas can be removed from the reaction product by an intermediate condensation or distillation.

Adjuvants, such as CCL; and/or tertiary butyl chloride, can be added directly to the reactor as by line 33, valve 34 and line 35, or can be added to the one of the feed components, such as the isoparaffin, or can be added to the paraffin-olefin mixture in the mixing vessel 17 or 18; however, in the event that the promoter can react with the microfilter 24, or cause corrosion in the pump 31, it is preferred that the adjuvant be added at some point after the pump as by line 65, valve 66 and line 67 (the high pressure side of the DP cell), thus the promoter becomes dispersed in the flowing nitrogen from the high pressure side of the DP cell and passes into the reactor below the liquid level and bubbles up through the reactor contents.

ILLUSTRATIVE EXAMPLES In the following examples, Example I shows the preparation of a preferred embodiment of the Dy catalyst and Example II shows the incorporation therewith of a Pt-hydrogenation catalyst. Example III shows continuous paraffin-olefin alkylation with the Dy zeolite catalyst of Example I.

Example I About 500 g. of NaY zeolite was exchanged, filtered and washed for 16 cycles with aqueous NH CI utilizing the procedures disclosed in the aforementioned U.S. application, Ser. No. 581,129. The resulting NH Y zeolite was similarly exchanged for 16 cycles with aqueous dysprosium nitrate. The aqueous dysprosium nitrate contained 11.8 g./l. Dy(NO .SH O (molecular weight, 438.6) and had a pH in the range of 4.3-4.7 (in making such solutions, the pH is commonly adjusted to about 4.5, if necessary, with HN or NH OH).

The resulting Dy-exchanged NH -exchanged zeolite was washed free of nitrate and unexchanged dysprosium ions, with distilled water, and dried in an oven at about 120 to produce a DyNH Y zeolite. The zeolite was activated by heating slowly in a stream of helium to 400 C. to remove water and decompose the bulk of any remaining ammonium ions. This activation utilized the procedures disclosed in the aforementioned U.S. Ser. No. 581,129. The DyNH Y zeolite before activation had the following analysis:

0.95 wt. percent Na 14.30 wt. percent Dy 1.24 wt. percent N 20.65% loss on ignition The base sodium zeolite, before exchange, analyzed 9.47% Na; 16.32% A1 0 47.87% SiO and had a 25.38% loss on ignition. The resulting substantially anhydrous DyHY zeolite was crystalline and capable of adsorbing benzene. The weight loss upon ignition analysis at 1800 C., of the activated zeolite was 3.72%.

Example II A solution of Pt(NH Cl in water was added dropwise with stirring to a dilute aqueous suspension of the dried DyNH Y zeolite of Example I at 55 C. The amount of Pt(NH Cl used was equivalent to 0.5% Pt in the activated catalyst. After the Pt salt addition was complete (about 1 hr.), the solution was stirred at 55 C. for 30 minutes, filtered, and the catalyst combination washed with distilled water until the washings were free of chloride ion. The catalyst combination was dried at 120 C., heated to 400 C. in a stream of dry air and then reduced at 400 C. in the reactor in a flowing stream of H for one hour. This catalyst combination is useful for the catalytic conversion of s-OHP at 200 C. under the conditions taught in US. 3,396,203 and is also useful for the conversion of s-OHP when there is also present in the reaction zone from 5-75 p.s.i. of hydrogen.

Example III In the apparatus of FIGS. 1, 2 and 3, a liquid mixture of 36 volumes of isobutane and 1 volume of butene-2 was fed, from the mixing section, and continuously contacted, in the reactor, in the presence of a halide adjuvant, with the helium activated DyHY catalyst of Example I. The liquid level in the reactor was maintained constant by the DP cell so that the concentration of the catalyst in the reaction mixture was about 10 weight percent. The total initial reactor charge was 2500 cc. at 25 C. and the total internal volume of the reactor was 3700 cc.

The feed rate and the rate of withdrawal of catalystfree reaction mixture were continuous and controlled, such that the mean retention time was 0.3 hour per gram of hydrocarbon per gram of catalyst. The catalyst was maintained primarily in suspension by continuous stirring.

The halide adjuvant was 0.14 g. t-butyl chloride and 0.23 g. CCL; per gram of olefin charged. The temperature in the reactor was maintained in the range of -100 C. and the pressure Was maintained at 25-50 p.s.i. above autogenous at the temperature (the hydrocarbon being maintained primarily in the liquid phase). The nitrogen flow in the high pressure side of the DP cell was 0.075 liter minute and in the low pressure side 0.350 l./m. (or a total flow of 0.425 l./m. of N Anhydrous HCl was added to the vapor space at an average rate equivalent to 0.25 mole HCl/mole of olefin charged. The HCl concentration was maintained in the vapor space in an intermittent manner by adding HCl for 30 minutes at a rate of about 0.1 l./m., then no further HCl was introduced for minutes. Such cycles of 30 minutes of HCl introduction in every minute period were continued throughout the run.

The C parafiin product recovered (continuously) for the first 39 hours of the run, from the so-withdrawn catalyst-free reaction mixture ranged from -236 wt. percent (most of the run being 200-210) based on the weight of olefin charged. The proportion of 2,2,4-trimethylpentane ranged from 23-23.6 mole percent of the trimethylpentane fraction, and relatively constant (most of the run being 25-29% indicating a continuing self-alkylgation function of the catalyst. The total trimethylpentane content of the C fraction of the liquid product was also relatively constant at about 88 to 91 mole percent of the C fraction of the continuously produced alkylate. C unsaturates in the liquid product ranged from 0.1 to 0.4 wt. percent. The total C alkylate (60+%, C produced in the first 39 hours per gram of activated DyHY catalyst was 9.95 g. For purposes of evaluation of catalysts in such a continuous alkylation process, the total yield is computed for the length of time that the run produced C liquid alkylate containing at least 60% of C paraflius, in the present application such a total or integrated yield is referred to as total C alkylate (60%+, C8.

Alkylate yield maxima as high as 230 to 265 wt. percent olefin charged can be obtained in runs similar to that described above. Although theoretical yield is 204 wt. percent, self-alkylation and cracking of isobutane-derived products can increase the observed yield.

I claim:

1. A composition comprising an acidic Dy aluminosilicate zeolite catalyst and from 0.05 to 5 percent of a hydrogenation catalyst, said Dy alumino-silicate containing less than one alkali metal cation and at least one divlent or trivalent metal, metal oxide or metal hydroxide for every 12 atoms of aluminum in said alumino-silicate 1 l zeolite framework, said framework having an Al/Si atomic ratio from 065-02, and wherein at least 25% of the electronegativity associated with the alumino-silicate framework is satisfied by at least one member selected from the group consisting of cations of dysprosium, of its oxides and of its hydroxides.

2. A composition according to claim 1 wherein said hydrogenation catalyst is selected from the group consisting of platinum, palladium, nickel, nickel oxide, nickel sulfide, molybdenum oxide, molybdenum sulfide, cobalt oxide, palladium oxide and mixtures thereof.

3. A composition according to claim 2 wherein said hydrogenation catalyst is physically admixed with the acidic alumino-silicate.

4. A composition according to claim 2 wherein said hydrogenation catalyst has been incorporated into the alumino-silicate by salt impregnation.

5. A composition according to claim 2 wherein said acidic Dy alumino-silicate catalyst is crystalline, and is capable of adsorbing benzene and wherein said hydrogenation catalyst has been introduced into the zeolite by ion exchange and wherein said Dy zeolite has an atomic ratio Al/H O in the range of 025-2.

6. A composition according to claim 5 wherein the hydrogenation catalyst is in a reduced form and wherein at least 40% of the electronegativity associated with the alumino-silicate framework is satisfied by at least one member selected from the class consisting of cations of dysprosium, its oxides and its hydroxides.

7. A hydrocarbon conversion process comprising contacting a hydrocarbonaceous feed in a conversion zone at an elevated conversion temperature with a combination of a Dy alumino-silicate zeolite catalyst and a hydrogenation catalyst and recovering a hydrocarbon conversion product having an average molecular weight no greater than that of said first-named hydrocarbonaceous feed, said Dy alumino-silicate containing less than one alkali metal cation and at least one divlent or trivalent metal, metal oxide or metal hydroxide for every 12 atoms of aluminum in said alumino-silicate zeolite framework, said framework an Al/ Si atomic ratio from 0.650.2 and wherein at least of the electronegativity associated with said framework is satisfied by at least one member selected from the group consisting of cations of dysprosium its oxides and its hydroxides.

8. A hydrocarbon conversion process comprising contacting a hydrocarbonaceous feed in a conversion zone at an elevated conversion temperature with a Dy aluminosilicate zeolite catalyst and recovering a hydrocarbon conversion product having an average molecular weight no greater than that of said first-named hydrocarbonaceous feed, said Dy alumino-silicate containing less than one alkali metal cation and at least one divalent or trivalent metal, metal oxide or metal hydroxide for every 12 atoms of aluminum in said alumino-silicate zeolite framework, wherein at least 25 of the electronegativity associated with said framework is satisfied by at least one member selected from the group consisting of dysprosium, its oxides and hydroxides, and wherein there is present in the conversion zone from 5-5000 p.s.i. hydrogen.

9. Process according to claim 7 wherein said feed comprises a polycyclic aromatic hydrocarbon.

10. Process according to claim 9 wherein said feed comprises s-OHP or s-OHA or a mixture thereof.

11. Process according to claim 10- wherein said catalyst comprises an acidic alumino-silicate zeolite catalyst and from 0.05 to 5 percent of a hydrogenation catalyst.

12. Process according to claim 11 wherein said Dy alumino-silicate catalyst is crystalline, is capable of adsorbing benzene and wherein said hydrogenation catalyst has been introduced into the zeolite by ion-exchange and is selected from the group consisting of platinum, palladium, nickel, nickel oxide, nickel sulfide, molybdenum oxide, molybdenum sulfide, cobalt oxide, palladium oxide and mixtures thereof and wherein said Dy zeolite has an atomic ratio Al/H O in the range of 0.25-2.

13. Process according to claim 10 wherein said hydrogenation catalyst is in a reducedform.

14. A hydrocarbon conversion process comprising contacting a hydrocarbonaceous feed in a conversion Zone at an elevated conversion temperature with the catalyst composition of claim 1 and recovering an upgraded hydrocarbon conversion product.

15. Process according to claim 14 wherein there is present in the conversion zone from 5-5000 p.s.i. of hydrogen.

References Cited UNITED STATES PATENTS 3,140,253 7/ 1964 Plank et a1. 208 3,247,099 4/ 1966 Oleck et a1. 208138 3,251,902 5/1966 Garwood ct a1. 260683.64 3,301,917 l/l967 Wise 260683.65 3,396,203 8/1968 Bushick 260668 DELBERT E. GANTZ, Primary Examiner C. R. DAVIS, Assistant Examiner U.S. Cl. X.R. 208120, 138; 252-455; 260-6832, 683.64 

